Low-carbon hydrogen process

Abstract

A process for producing hydrogen is described, comprising: (i) subjecting a gaseous mixture comprising hydrocarbons and steam to steam reforming in a gas-heated reformer or an adiabatic pre-reformer, followed by autothermal reforming with an oxygen-rich gas in an autothermal reformer to produce a reformed gas mixture; (ii) increasing the hydrogen content of the reformed gas mixture by subjecting the reformed gas mixture to one or more water-gas shift stages in a water-gas shift unit to provide a hydrogen-rich reformed gas; (iii) passing the hydrogen-rich reformed gas and the oxygen-rich gas through an oxidation unit comprising an oxidation catalyst that converts carbon monoxide present in the hydrogen-rich reformed gas to carbon dioxide to form a carbon dioxide-rich gas mixture; (iv) cooling the carbon dioxide-rich gas mixture and separating condensed water therefrom; and (v) passing the carbon dioxide-rich gas mixture to a carbon dioxide separation unit to provide a carbon dioxide gas stream and a hydrogen product gas stream.

Description

[Technical field]

[0001] The present invention relates to a process for converting hydrocarbons to hydrogen while minimizing the production and emissions of carbon dioxide.

[0002] Processes to produce hydrogen are well known and generally include combustion steam methane reformers combined with water gas shift and carbon dioxide (CO2) removal. Such processes produce significant volumes of carbon dioxide in the flue gas at pressures that are not suitable for efficient CO2 capture. A general goal is to increase the rate of progress towards net zero targets. In the meantime, there is a need for hydrogen production processes that produce lower levels of carbon dioxide effluent and allow for more efficient CO2 capture.

[0003] Supporting this, blue hydrogen or low carbon hydrogen processes are under development. A process for low carbon hydrogen is disclosed in a paper entitled "Clean Hydrogen. Part 1: Hydrogen from Natural Gas through Cost Effective CO2 Capture" published in The Chemical Engineer (March 15, 2019). The process disclosed in the paper included the steps of desulfurization, saturation, reforming in a gas-heated reformer and an oxygen-fed autothermal reformer, isothermal temperature shift, cooling, condensate removal, and pressure swing adsorption. The percentage of captured CO2 was 95.4% in the LCH process.

[0004] The present inventors have developed an improved process that reduces carbon dioxide emissions.

[0005] Accordingly, the present invention provides a process for producing hydrogen, comprising the steps of: (i) subjecting a gaseous mixture comprising hydrocarbons and steam to steam reforming in a gas-heated reformer or an adiabatic pre-reformer, followed by autothermal reforming with an oxygen-enriched gas in an autothermal reformer to produce a reformed gas mixture; (ii) increasing the hydrogen content of the reformed gas mixture by subjecting the reformed gas mixture to one or more water-gas shift stages in a water-gas shift unit to provide a hydrogen-enriched reformed gas; (iii) passing the hydrogen-rich reformed gas and the oxygen-rich gas through an oxidation unit containing an oxidation catalyst that converts carbon monoxide present in the hydrogen-rich reformed gas to carbon dioxide to form a carbon dioxide-rich gas mixture; (iv) cooling the carbon dioxide-enriched gas mixture and separating condensed water therefrom; (v) passing the carbon dioxide enriched gas mixture to a carbon dioxide separation unit to provide a carbon dioxide gas stream and a hydrogen product gas stream.

[0006] Lower CO2 emissions can be achieved by converting carbon monoxide to carbon dioxide using an oxidation catalyst in the oxidation unit, followed by removal by a carbon dioxide separation unit. Furthermore, a fuel-grade hydrogen product gas stream can be obtained without the use of expensive PSA units. Furthermore, a tunable portion of the product hydrogen can be combusted to provide process heating and power loads, thus greatly enhancing the flexibility of plant operation.

[0007] The present invention can provide all product hydrogen at high pressure, which paves the way for more exergy-efficient methods of utilizing hydrogen as a fuel or as a feedstock for downstream processes.

[0008] The present invention uses a gaseous mixture that contains hydrocarbons.

[0009] The gaseous mixture may comprise any gaseous or low boiling hydrocarbon such as natural gas, associated gas, LPG, petroleum distillates, diesel, naphtha or mixtures thereof, or a hydrocarbon-containing off-gas from a chemical process such as refinery off-gas. The gaseous mixture comprises methane, preferably associated gas, or natural gas, containing a significant proportion of methane, for example more than 50 v / v %. Natural gas is particularly preferred. The hydrocarbons may be compressed to a pressure in the range of 10 bar abs to 100 bar abs. The pressure of the hydrocarbons may usefully govern the pressure throughout the process. The operating pressure is preferably in the range of 15 bar abs to 50 bar abs, more preferably 25 bar abs to 50 bar abs, as this improves performance from the process.

[0010] If the hydrocarbons contain sulfur compounds before or preferably after compression, they may be subjected to desulfurization, including hydrodesulphurization using CoMo or NiMo catalysts, and adsorption of hydrogen sulfide using a suitable hydrogen sulfide adsorbent, such as a zinc oxide adsorbent. Ultra-purified adsorbents may be usefully used downstream of the hydrogen sulfide adsorbent to further protect the steam reforming catalyst. Suitable ultra-purified adsorbents may include copper-zinc oxide / alumina materials and copper-nickel-zinc oxide / alumina materials. It is preferred to add hydrogen to the compressed hydrocarbons to facilitate hydrodesulphurization and / or to reduce the risk of carbon laydown in the reforming process. The amount of hydrogen in the resulting mixed gas stream may be in the range of 1% to 20% by volume on a dry gas basis, but is preferably in the range of 1% to 10% by volume, more preferably in the range of 1% to 5% by volume. In a preferred embodiment, a portion of the hydrogen product stream may be mixed with the compressed hydrocarbons. Hydrogen may be combined with the hydrocarbons upstream and / or downstream of any hydrodesulphurization stage.

[0011] If the hydrocarbons contain other contaminants such as chlorides or heavy metal contaminants, these may be removed using conventional adsorbents upstream or downstream of any desulfurization prior to upgrading. Adsorbents suitable for chloride removal are well known and include alkalized alumina materials. Similarly, adsorbents for heavy metals such as mercury or arsenic are known and include copper sulphide materials.

[0012] The hydrocarbons may be preheated. The hydrocarbons may be conveniently preheated after compression and before desulfurization. The process provides a variety of hot gas sources that can be used for this application. However, in a preferred embodiment, the hydrocarbons are heated by passing them through a combustion heater that is fueled by a portion of the hydrogen product stream.

[0013] The hydrocarbons are mixed with steam. Steam introduction may be by direct injection of steam and / or by saturation of the hydrocarbons by contact with a heated water stream. In some arrangements, a gaseous mixture containing hydrocarbons and steam may be formed by directly mixing the hydrocarbons with steam, for example steam generated in one or more fired heaters, and / or by cooling the reformed gas mixture with water. In other arrangements, the hydrocarbons may be saturated in a saturator supplied with hot water to form a saturated gas mixture. The water may include one or more condensate streams produced in the process. The steam content of the saturated gas mixture may be increased by direct addition of steam, if desired.

[0014] In arrangements using gas heater reformers or autothermal reformers, the amount of steam introduced is desirably sufficient to provide a steam to carbon ratio at the inlet to the reforming unit operation of at least 2.5:1, i.e., at least 2.5 moles of steam per gram atom of hydrocarbon in the gaseous mixture. For efficient utilization of energy in the process, the steam to carbon ratio may be high, which maximizes hydrogen production. The steam to carbon ratio may be usefully up to about 5:1.

[0015] In an arrangement using an adiabatic pre-reformer and an autothermal reformer, it is desirable that the amount of steam introduced is sufficient to provide a steam-to-carbon ratio (defined as the ratio of steam to hydrocarbon carbon at the inlet to the reforming unit operation) of at least 0.9:1, i.e., at least 0.9 moles of steam per gram of hydrocarbon carbon atom in the gaseous mixture, with a preferred range of 0.9:1 to 5:1. If the steam-to-carbon ratio at the inlet to the reforming unit operation is in the range of 0.9:1 to less than 2.4:1, additional steam needs to be added to the reformed gas upstream of the water-gas shift stage. Operating the reforming section with a steam-to-carbon ratio in the range of 0.9:1 to less than 2.4:1 has the advantage that the heating requirements and oxygen demand for the reforming stage are reduced. If the steam-to-carbon ratio is in the range of 2.4:1 to 5:1, no further steam addition upstream of the water-gas shift unit is required, which may be useful in situations where steam addition to the reformed gas is impractical.

[0016] The gaseous mixture containing the hydrocarbons and steam is then desirably preheated prior to reforming. In a preferred embodiment, the gaseous mixture is heated by passing it through a combustion heater fueled by a portion of the hydrogen product stream, in particular the same combustion heater used to preheat the hydrocarbons. Desirably, the mixed stream is heated to between 400°C and 500°C, preferably between 420°C and 460°C.

[0017] In some arrangements, the invention includes an adiabatic pre-reforming stage upstream of the autothermal reformer. In these arrangements, the gaseous mixture containing the hydrocarbons and steam is subjected to a step of adiabatic steam reforming in a pre-reforming vessel containing a fixed bed of pre-reforming catalyst. In such a process, the gaseous mixture containing the hydrocarbons and steam is adiabatically passed through a bed of steam reforming catalyst, usually with a high nickel content, e.g., greater than 40 wt.%, at an inlet temperature typically in the range of 400°C to 650°C. During such an adiabatic pre-reforming step, any hydrocarbons higher than methane are reacted with steam to obtain a mixture of methane, carbon oxides, and hydrogen. The use of such an adiabatic steam reforming step, commonly referred to as pre-reforming, may be desirable to ensure that the feed to the autothermal reformer does not contain any hydrocarbons higher than methane, and also contains some hydrogen.

[0018] In another arrangement, a gaseous mixture comprising hydrocarbons and steam is subjected to steam reforming in a gas-heated reformer.

[0019] In one type of gas-heated reformer, the catalyst is disposed in tubes that extend between a pair of tube sheets through a heat exchange zone. The reactants are fed into the zone above the upper tube sheet and pass through the tubes into the zone below the lower tube sheet. The heating medium passes through the zone between the two tube sheets. The heating medium is typically hot reformed gas recovered from an autothermal reformer. Gas-heated reformers of this type are described in GB 1578270 and WO 97 / 05947. Another type of gas-heated reformer that may be used is a dual-tube gas-heated reformer as described in U.S. Pat. No. 4,910,228, in which the reformer tubes each include an outer tube having a closed end and an inner tube concentrically disposed within the outer tube, and communicate with the annular space between the inner and outer tubes at the closed end of the outer tube with a steam reforming catalyst disposed within the annular space. The outer surface of the outer tube is heated by a heating medium. A reaction mixture is fed to the end of the outer tube remote from the closed end such that the mixture passes through the annular space, undergoes steam reforming, and then passes through the inner tube.

[0020] The compressed preheated gaseous mixture containing hydrocarbons and steam passes through catalyst-filled tubes in the gas-heated reformer. During the passage through the reforming catalyst, the endothermic steam reforming reaction takes place with the heat required for the reaction being provided by the hot reformed gas recovered from the autothermal reformer, flowing past the outer surface of the tube. The steam reforming catalyst used in the gas-heated reformer may comprise nickel supported on a particulate refractory support such as rings or porous pellets of calcium aluminate, magnesium aluminate, alumina, titania, and zirconia. Alternatively, a combination of nickel and a precious metal such as ruthenium may be used. Instead of or in addition to a particulate steam reforming catalyst, the steam reforming catalyst may comprise one or more structured catalyst units, which may be in the form of a metal or ceramic monolith or folded metal structure on which layers of nickel and / or precious metal steam reforming catalyst are deposited. Such structured catalysts are described, for example, in WO 2012 / 103432 A1 and WO 2013151885 A1. The temperature of the autothermal reforming gas used to heat the gas-heated reformer is preferably sufficient such that the gas undergoing steam reforming leaves the catalyst tubes at a temperature in the range of 600°C to 850°C, preferably 650°C to 750°C, more preferably 680°C to 720°C.

[0021] In the present invention, the pre-reformed gas or steam reformed gas containing methane, hydrogen, steam, and carbon oxides is preferably fed directly to the autothermal reformer for autothermal reforming without any dilution or heat exchange. The pre-reformer or gas heated reformer and the autothermal reformer operate in series.

[0022] The autothermal reformer may comprise a vertically arranged elongated vessel with a burner located at the top of the reformer to which the pre-reformed or steam reformed gas and the oxygen-enriched gas are fed, a combustion zone below the burner through which the flame extends, and a fixed bed of particulate steam reforming catalyst located below the combustion zone. Thus, in autothermal reforming, heat for the endothermic steam reforming reaction is provided by combustion of some of the hydrocarbons in the feed gas. The pre-reformed or steam reformed gas is typically fed to the top of the reformer, and the oxygen-enriched gas is fed to the burner, with mixing and combustion occurring downstream of the burner to produce a heated gas mixture whose composition is equilibrated as it passes through the steam reforming catalyst. The autothermal steam reforming catalyst may comprise nickel supported on a refractory support such as rings or pellets of calcium aluminate, magnesium aluminate, alumina, titania, zirconia, and the like. In a preferred embodiment, the autothermal steam reforming catalyst comprises a layer of catalyst comprising Ni and / or Ru on zirconia over a bed of Ni on alumina catalyst to reduce catalyst support volatilization that can result in reduced performance of the autothermal reformer.

[0023] The oxygen-enriched gas may contain at least 50% O2 by volume, or may be an oxygen-enriched air mixture. However, in the present invention, the oxygen-enriched gas preferably comprises at least 90% O2 by volume, more preferably at least 95% O2 by volume, most preferably at least 98% O2 by volume, or at least 99% O2 by volume, such as a pure oxygen gas stream obtainable using a vacuum pressure swing adsorption (VPSA) or air separation unit (ASU). The ASU may be electrically powered, preferably powered using renewable electricity to further improve the efficiency of the process and minimize CO2 emissions.

[0024] The amount of oxygen-enriched gas added is preferably such that between 40 and 60 moles of oxygen are added per 100 moles of carbon in the hydrocarbon fed to the process. The amount of oxygen added is preferably such that the reformed gas leaves the catalyst in the autothermal reformer at a temperature in the range of 800° C. to 1100° C., more preferably 900° C. to 1100° C., most preferably 970° C. to 1070° C. In a preferred embodiment, a small steam purge may be added to the oxygen-enriched gas to prevent backflow in the event of a plant trip.

[0025] In an arrangement including a pre-reformer and an autothermal reformer, the autothermal reformed gas is cooled, for example in a waste heat boiler, against water to generate steam, which can be used for heating and / or power generation in a steam turbine.

[0026] In an arrangement with a gas-heated reformer and an autothermal reformer, the reformed gas produced by the autothermal reformer is used to provide the heat required for the steam reforming step by using it as hot gas passing through tubes in the gas-heated reformer. During this heat exchange, the reformed gas transfers heat to the gas being steam reformed, cooling it. Preferably, the reformed gas leaves the gas-heated reformer at a temperature somewhat higher than the temperature at which the gaseous mixture comprising the hydrocarbon and steam mixture is fed to the gas-heated reformer, although it cools by a few hundred degrees Celsius. Preferably, the reformed gas leaves the gas-heated reformer at a temperature in the range of 450° C. to 650° C., more preferably 450° C. to 580° C. After leaving the gas-heated reformer, the reformed gas is further cooled in one or more steps of heat exchange. The heat recovered during this cooling may be employed to preheat the reactants and / or to heat the water used to provide the steam employed in the steam reforming step. In some arrangements, the reformed gas mixture exiting the shell side of the gas-heated reformer may be used to heat water that is fed to the saturator.

[0027] Heat recovered from the autothermal reformed gas may additionally or alternatively be used in the carbon dioxide separation step.

[0028] The reformed gas contains hydrogen, carbon monoxide, carbon dioxide, steam, and small amounts of unreacted methane, and may contain small amounts of inert gases, such as nitrogen and argon. Preferably, the hydrogen content of the reformed gas ranges from 30% to 45% by volume on a wet gas basis, and the carbon monoxide content ranges from 5% to 15% by volume on a wet gas basis. In the present invention, the hydrogen content of the reformed gas mixture is increased by subjecting the reformed gas mixture to one or more water-gas shift stages, thereby producing a hydrogen-enriched reformed gas and simultaneously converting the carbon monoxide in the reformed gas to carbon dioxide. The reaction can be illustrated as follows:

[0029] [ka]

[0030] Steam is present in the reformed gas, but if desired, additional steam may be added prior to one or more water gas shift stages, for example by direct addition to the reformed gas.

[0031] The reformed gas may be subjected to one or more water-gas shift stages in a water-gas shift unit to form a hydrogen-enriched reformed gas stream or a "shifted" gas stream. The one or more water-gas shift stages may include a high temperature shift, a mid-temperature shift, an isothermal shift, and a low temperature shift stage.

[0032] The high temperature shift is operated adiabatically in a shift vessel having an inlet temperature in the range of 300° C. to 400° C., preferably 320° C. to 360° C., over a bed of reduced iron catalyst such as chromia promoted magnetite. Alternatively, a promoted zinc aluminate catalyst may be used.

[0033] The medium temperature shift stage and the low temperature shift stage may be carried out using a shift vessel containing a supported copper catalyst, in particular a copper / zinc oxide / alumina composition. In the low temperature shift, gas containing carbon monoxide (preferably not more than 6% CO by volume on a dry basis) and steam (at a steam to total dry gas molar ratio in the range of 0.3 to 1.5) may be passed over the catalyst in an adiabatic fixed bed with an outlet temperature in the range of 200°C to 300°C. The outlet carbon monoxide content may be in the range of 0.1% to 1.5% by volume on a dry basis, in particular less than 0.5% by volume, if additional steam is added. Alternatively, in the medium temperature shift, gas containing carbon monoxide and steam may be fed to the catalyst at an inlet temperature in the range of 200°C to 240°C, although the inlet temperature may be as high as 280°C. The outlet temperature may be up to 300°C, but may be as high as 360°C.

[0034] The reformed gas is preferably subjected to an isothermal water gas shift stage in a cold shift vessel, optionally followed by one or more adiabatic medium or low temperature water gas shift stages in a non-cooled vessel as described above, although one or more adiabatic water gas shift stages, such as a high temperature shift stage, optionally followed by a low temperature shift stage, may be employed. The use of an isothermal shift stage, i.e. heat exchange in a shift converter, such that an exothermic reaction takes place on a catalyst bed in contact with a heat exchange surface that removes heat, offers the possibility of using the reformed gas stream in a very efficient manner. The term "isothermal" is used to describe a cold shift converter, while the temperature of the hydrogen-rich reformed gas stream at the outlet of the isothermal shift converter may be 1°C to 25°C higher than the inlet temperature, with the possibility of a slight increase in the temperature of the gas between the inlet and the outlet. The coolant may conveniently be water under pressure such that partial or complete boiling occurs. The water may be in tubes surrounded by the catalyst or vice versa. The resulting steam may be used, for example, to drive a turbine, for example for electricity, or to provide process steam for feeding the process. In a preferred embodiment, the steam produced by the isothermal shift stage is used to supplement the addition of steam to a gaseous mixture comprising hydrocarbons and steam upstream of a pre-reformer or gas-heated reformer.

[0035] The addition of an adiabatic mid-temperature or low-temperature shift stage downstream of the isothermal shift stage offers the potential to increase the CO2 capture efficiency from the process to 98% or more, however, the inventors have found that superior efficiency can be provided by a single isothermal shift converter.

[0036] The hydrogen-rich reformed gas includes steam. In some embodiments, following one or more water-gas shift stages, it may be desirable to cool the hydrogen-rich reformed gas to a temperature below the dew point so that at least a portion of the steam condenses. The liquid water condensate may then be separated using one or more gas-liquid separators, which may have one or more further cooling stages between them to form a dehydrated hydrogen-rich reformed gas. Any cooling agent may be used. Preferably, the cooling of the hydrogen-rich reformed gas stream is performed in heat exchange with the process condensate. As a result, a heated water stream is formed that may be used to provide some or all of the steam required for reforming. Thus, in some arrangements, the condensate recovered from the hydrogen-rich reformed gas is used to obtain at least a portion of the steam in the gas mixture that is fed to the steam reforming step. Since the condensate may contain ammonia, methanol, hydrogen cyanide, and CO2, returning the condensate to steam provides a useful process for returning hydrogen and carbon to the process. One or more stages of cooling and condensate recovery may be included upstream of the oxidation unit, but in a preferred arrangement there is a single stage of cooling and condensate recovery, such that the oxidation unit is located after the first gas-liquid separator.

[0037] In other embodiments, the oxidation unit may be located downstream of the water-gas shift unit without intermediate cooling and condensate separation, in which case the oxidation unit is fed directly with the steam-containing hydrogen-rich reformate gas.

[0038] In the present invention, the hydrogen-rich reformed gas is fed to an oxidation unit, after cooling and separation of condensate if necessary. In the oxidation unit, the carbon monoxide present in the hydrogen-rich reformed gas is oxidized to carbon dioxide. An oxygen-rich gas is added to the hydrogen-rich reformed gas. Some hydrogen is also oxidized to form water. Therefore, the oxidation catalyst is preferably a CO-selective oxidation catalyst to minimize hydrogen loss.

[0039] The reaction can be shown as follows: 2 H2+O2→2 H2O 2CO+O2→2CO2

[0040] The oxidation catalyst is preferably a supported precious metal catalyst. For example, the catalyst may comprise one or more of Pt, Pd, Rh, Ir or Ru, desirably on an oxide support such as alumina, titania, zirconia or silica. The amount of precious metal may range from 0.1% to 5% by weight. The oxidation catalyst may be in the form of pellets or extrudates, foams, monoliths or coatings on inert supports. A suitable precious metal oxidation catalyst for CO oxidation preferably comprises platinum on alumina promoted with an oxide of a metal selected from the group consisting of manganese, iron, cobalt, copper, nickel and mixtures thereof. A particularly suitable catalyst is an alumina supported platinum catalyst promoted with iron oxide and / or copper oxide. The loading of platinum on the particulate support material should be in the range of about 1% to 5% by weight, preferably about 1% to 3% by weight. The loading of copper, if present, should be about 2% to 12% by weight, preferably 4% to 8% by weight. The iron loading, if present, is desirably about 0.1% to 2% by weight, preferably about 0.2% to 1% by weight. Such catalysts are described in US Patent Application Publication No. 2006276332(A1), US Patent No. 6,559,094 and US Patent No. 3,088,919. Alternatively, the oxidation catalyst may comprise supported copper oxide. For example, CN102407123A discloses CuO supported on ceria as a preferential CO oxidation catalyst. The selectivity of the oxidation catalyst for CO oxidation is preferably at least 50%.

[0041] The oxidation reaction is exothermic and the reaction can be carried out adiabatically in a fixed bed oxidation vessel. Flow through the bed can be axial and / or radial. Inlet temperatures in such an arrangement can range from 20°C to 200°C. The exotherm in the bed is desirably kept below 75°C. Lower temperatures in the catalyst bed generally favor higher selectivity to CO oxidation and are preferred. It is therefore preferred to place the adiabatic oxidation vessel downstream of one or more stages of cooling and condensate recovery. In this way higher selectivity can be achieved which limits the temperature rise across the adiabatic catalyst bed and gives longer catalyst life.

[0042] In some embodiments, it is preferred to operate the oxidation unit with staged oxygen addition. In such embodiments, the oxidation unit may include two or more adiabatic acidification reactors in series. A first oxygen-containing gas stream can be fed to the first oxidation reactor to perform the first stage of selective CO oxidation. The effluent from the first oxidation reactor can be cooled, for example, by indirect heat exchange with cooling water or a process stream, to form a first cooled effluent stream. A second oxygen-containing gas stream can be added to the first cooled effluent stream, and the resulting mixture can be fed to a second oxidation reactor to perform the second stage of selective CO oxidation. This sequence can be repeated with additional reaction, cooling, and oxygen addition stages.

[0043] Alternatively, the oxidation step may be operated with cooling of the catalyst bed by passing a gas or liquid coolant, such as a suitable process stream, preferably steam or a boiling water-steam mixture, through one or more tubes disposed within the catalyst bed. Alternatively, the catalyst may be in a tube surrounded by boiling water. If a cooled oxidation reactor is used, it is possible to place the oxidation unit directly downstream of the water-gas shift unit, especially if the water-gas shift unit includes an isothermal shift vessel and there is no cooling stage in between. In this arrangement, the gas mixture fed to the oxidation unit is a hydrogen-rich reformed gas containing steam, and the inlet temperature of the oxidation unit is close to the outlet temperature from the water-gas shift unit, which may be in the range of 200°C to 320°C. If the coolant is water under pressure such that partial or complete boiling occurs, the water pressure is preferably the same as in the isothermal shift converter, so that a single steam drum may be shared by the isothermal water-gas shift and oxidation units. Alternatively, the water pressure in the cooled oxidation unit may be different, preferably lower, than the water pressure in the isothermal water-gas shift unit.

[0044] In embodiments where an isothermal water-gas shift converter is not used, a cooled oxidation unit may also be used. In such embodiments, the water-gas shift unit may include one or more adiabatic reactors. The water pressure in the cooled oxidation unit may range from atmospheric to 50 bar.

[0045] Further, to limit peak temperatures in the cooled oxidation unit, the oxidation unit may contain two or more cooled catalyst beds in series, and a portion of the oxygen-rich gas may be added stepwise between successive catalyst beds. The two or more cooled catalyst beds may be contained within the same pressure vessel or in different pressure vessels.

[0046] In the present invention, the oxidation unit is fed with oxygen-rich gas, preferably a portion of the same oxygen-rich gas fed to the autothermal reformer. Thus, the oxygen-rich gas fed to the oxidation unit preferably comprises at least 90% by volume O2, more preferably at least 95% by volume O2, most preferably at least 98% by volume O2, or at least 99% by volume O2, e.g. a pure oxygen gas stream. In some arrangements, the oxygen may be provided by electrolysis, by using a renewable electricity source, or by electricity generated by a turbine powered from steam provided from the process, e.g., from the combustion of a portion of the hydrogen product, or from the combustion of a portion of the hydrogen product. To ensure a high conversion of residual carbon monoxide, oxygen may be added in stoichiometric excess, but too high an excess may cause undesirable side reactions. As a result, it is preferred that the oxygen in the gas mixture is in less than 100% excess over stoichiometry.

[0047] The oxidation unit produces a carbon dioxide enriched gas mixture. Following the oxidation step, the carbon dioxide enriched gas mixture is cooled below the dew point to cause condensation of steam present in the gas mixture.

[0048] Cooling may be accomplished in one or more stages of heat exchange with water, air, or a combination thereof. In a preferred embodiment, cooling may also be accomplished in heat exchange with one or more liquids in a CO2 separation unit to improve process efficiency.

[0049] The cooled gas mixture may then be fed to a first gas-liquid separator to separate the gas mixture from the condensate to form a dehydrated carbon dioxide-rich gas stream. If desired, the separated gas may be further cooled with water and / or air and fed to a second separator, and then further cooled with water and / or air and fed to a third separator. Some or all of the condensate may be used to produce steam for the steam reforming and / or water gas shift steps.

[0050] Carbon dioxide is separated from the resulting dehydrated carbon dioxide rich gas stream to produce a hydrogen product stream.

[0051] The carbon dioxide separation step may be carried out using a physical or reactive scrubbing system, preferably a reactive scrubbing system, in particular an amine scrubbing system. Carbon dioxide may be separated by an acid gas recovery (AGR) process. In the AGR process, the carbon dioxide-rich gas stream is contacted with a suitable sorbent liquid, such as a stream of an amine, in particular a methyl diethanolamine (MDEA) solution, so that the carbon dioxide is adsorbed by the liquid to obtain a loaded sorbent liquid and a gas stream with a reduced carbon dioxide content. The loaded absorbent liquid is then regenerated by heating and / or depressurization to desorb the carbon dioxide and obtain a regenerated absorbent liquid, which is then recycled to the carbon dioxide adsorption step. Alternatively, methanol or glycol may be used to capture the carbon dioxide in a similar manner to the amines. In one arrangement, at least a portion of this heating is heat exchanged with a hydrogen-rich reformed gas stream recovered from the water gas shift unit. In another arrangement, at least a portion of this heating is heat exchanged with a carbon dioxide-rich gas mixture recovered from the oxidation unit. If the carbon dioxide separation step is operated as a single pressure process, i.e. essentially the same pressure is employed in the adsorption and regeneration steps, only little recompression of the recycled carbon dioxide is required.

[0052] For example, carbon dioxide captured from AGR may be compressed and used to produce chemicals such as methanol, or sent to storage or sequestration, or used in enhanced oil recovery (EOR) processes. Compression may be achieved using an electrically driven compressor powered by renewable electricity. When the CO2 is compressed for storage, transportation, use in EOR processes, or conversion to other chemical products, it may be dried to prevent traces of liquid water from condensing. For example, the CO2 may be dried to a dew point of -10°C or lower by passing it through a bed of a suitable desiccant, such as zeolite, or by contacting it with glycol in a glycol drying unit.

[0053] Upon separating the carbon dioxide, the process provides a crude hydrogen product gas stream. The crude hydrogen stream may comprise 95% to 99% by volume hydrogen, the remainder comprising methane, carbon monoxide, carbon dioxide, and inert gases. The methane content may range from 0.25% to 1.5% by volume, preferably 0.25% to 0.5% by volume. The carbon monoxide content may be less than 100 ppmv or 50 ppmv, preferably less than 20 ppmv, more preferably less than 10 ppmv. The carbon dioxide content may range from 0.01% to 0.5% by volume, preferably 0.01% to 0.1% by volume. The remainder may consist of nitrogen and residual steam.

[0054] For example, if the hydrogen product is to be transported by pipeline, it is desirably pre-dried. The drying step can be carried out using conventional glycol dryers or molecular sieves, similar to the drying of CO2.

[0055] The hydrogen gas stream is sufficiently pure for many loads, but if desired, the hydrogen product gas stream may be passed through a purification unit to provide purified hydrogen gas and a fuel gas. If used, at least a portion of the fuel gas may be used in the process as a replacement for an external fuel source to minimize CO2 emissions from the process. The purification unit may suitably comprise a membrane system, a temperature swing adsorption system, or a pressure swing adsorption system. Such systems are commercially available. The purification unit may produce a pure hydrogen stream, preferably having a purity of greater than 99.5% by volume, more preferably greater than 99.9% by volume.

[0056] However, unlike prior art processes that require high purity hydrogen, in the present invention, a purification unit such as a PSA unit is not required, and therefore the process can operate without a purification unit such as a pressure swing absorption unit.

[0057] The hydrogen product gas, with or without purification in the purification unit, may be compressed and used in downstream power or heating processes, for example by using it as fuel in a gas turbine (GT) or by injection into a domestic or industrial network gas piping system. The hydrogen product may also be used in downstream chemical synthesis processes, optionally after further purification. Thus, the hydrogen product may be purified and used to produce ammonia by reaction with nitrogen in an ammonia synthesis unit. Alternatively, the hydrogen product may be purified and used with a carbon dioxide-containing gas to produce methanol in a methanol production unit. Alternatively, the hydrogen product may be purified and used with a carbon monoxide-containing gas to synthesize hydrocarbons in a Fischer-Tropsch production unit. Any known ammonia, methanol, or Fischer-Tropsch production technology may be used. Alternatively, the hydrogen may be used to upgrade hydrocarbons, for example by hydrotreating or hydrocracking of hydrocarbons in a hydrocarbon refinery, or in any other process where pure hydrogen may be used. Compression may again be achieved using an electrically driven compressor powered by renewable electricity.

[0058] A portion of the hydrogen product, with or without purification, can be compressed if necessary and recycled to the hydrocarbon feed as desired for desulfurization to reduce the potential for carbon formation in the pre-reformer or gas-heated reformer.

[0059] In some arrangements, the hydrogen can be burned directly in a gas turbine without the need for recompression. The turbine exhaust gas can be used to raise process steam to meet the heating requirements of the process.

[0060] If the application of the hydrogen is to generate electricity, most of the product hydrogen can be sent to a gas turbine and the excess heat can be used to generate further electricity, for example by generating high pressure steam and expanding it through a steam turbine, as is done in conventional combined cycle power plants. Intermediate pressure steam can be extracted from the appropriate stages of the steam turbine and sent to the process, thus improving the heat integration in the hydrogen plant and the power plant for better energy utilization. Heat recovery operations can also be completed at low temperatures from both the hydrogen and power plant by district heating.

[0061] In this way, a better use of the energy of the process can be achieved compared to prior art processes. The mere addition of an oxidation unit leads to a reduction of CO2 emissions by up to 60%. 3 For a 1000 kW / h hydrogen plant, this corresponds to about 30 tonnes of CO2 per day, or 10,000 tonnes per year. Omitting the PSA unit results in higher operational flexibility in addition to lower capital expenditure. Using a hydrogen-fired gas turbine in the process provides higher exergy efficiency, lower electricity import (or net electricity export). Providing heat integration between the hydrogen plant and the power plant also gives a higher exergy efficiency than if the plants were operating separately, since the medium pressure steam extracted from the steam turbine to the hydrogen plant is doing some work in the high pressure stage of the turbine to generate electricity. Furthermore, if low temperature heat recovery from the hydrogen and / or power plant to the heat network is also included, the energy efficiency of the entire system is further improved. [Brief description of the drawings]

[0062] The present invention will now be described with reference to the accompanying drawings. [Figure 1] 1 shows a flow sheet of one embodiment of the present invention. [Diagram 2] 2 shows a flow sheet of another embodiment of the present invention. [Diagram 3]4 shows a flow sheet of yet another embodiment of the present invention. [Figure 4] 4 shows a flow sheet of yet another embodiment of the present invention.

[0063] The drawings are schematic and those skilled in the art will understand that in a commercial plant, additional items of equipment may be required, such as reflux drums, pumps, vacuum pumps, temperature sensors, pressure sensors, pressure relief valves, control valves, flow controllers, level controllers, holding tanks, storage tanks, etc. Providing such accessories of equipment does not form a part of the present invention and is in accordance with conventional chemical engineering practice.

[0064] In FIG. 1, a natural gas stream 10 is processed in a purification and saturation unit 12 for desulfurization, including hydrodesulfurization over a hydrodesulfurization catalyst and adsorption of hydrogen sulfide on a hydrogen sulfide adsorbent, followed by saturation with steam using a saturator to form a saturated natural gas mixture 14. The steam ratio of the saturated natural gas mixture 14 is increased by the addition of saturated medium pressure steam via line 16 and superheated medium pressure steam via line 18. The resulting gaseous mixture 20 of steam and natural gas is sent to a synthesis gas production unit 22, which includes a gas-heated reformer (GHR) and an autothermal reformer (ATR) in series. The gaseous mixture 20 is subjected to a first step of steam reforming in the gas-heated reformer over a granular nickel steam reforming catalyst arranged in a number of external heating tubes, followed by autothermal reforming with an oxygen gas stream 24 in the autothermal reformer, which includes a nickel steam reforming catalyst bed. The reformed gas recovered from the autothermal reformer is used to heat the tubes of the gas-heated reformer, thereby producing a partially cooled reformed gas 26, which is fed to a water-gas shift unit 28, which includes an isothermal water-gas shift vessel containing a bed of water-gas shift catalyst cooled by boiling water under pressure. The exothermic water-gas shift reaction generates steam, which is recovered from the water-gas shift unit 28 via line 30. A hydrogen-rich gas mixture is recovered from the water-gas shift unit 28 and fed via line 32 to an oxidation unit 34, which includes an oxidation vessel containing a fixed bed of a noble metal CO oxidation catalyst. A further oxygen-rich gas stream is added to the oxidation vessel via line 36. Carbon monoxide is oxidized to form a carbon dioxide-rich gas mixture. In this embodiment, the CO oxidation catalyst in the oxidation vessel is cooled by boiling pressurized water, which is converted to steam by an exothermic oxidation reaction. Steam is recovered from the oxidation unit 34 via line 38. In this embodiment, the pressure of steam 38 is the same as the steam recovered via line 30, so that streams 38 and 30 are combined to form the auxiliary steam that is added to the saturated natural gas 14 via line 16. The carbon dioxide-rich gas recovered from the oxidation unit 34 via line 40 is subjected to heat recovery in a heat recovery unit 42 which, in a final step, cools the carbon dioxide-rich gas 40 below the dew point in heat exchange with cold water to condense the steam.The liquid condensate is separated from the gas in one or more gas-liquid separators (not shown) and fed to the saturator of unit 12. The dehydrated carbon dioxide-rich gas recovered from the one or more gas-liquid separators is fed via line 44 to a carbon dioxide removal unit 46 operating by reactive absorption, which removes carbon dioxide from the gas mixture, thereby producing a carbon dioxide stream 48 that is compressed and sent for storage or conversion to chemicals. By removing carbon dioxide from the feed gas 44, the carbon dioxide removal unit produces a hydrogen product gas stream that is recovered from the carbon dioxide removal unit via line 50. In this embodiment, the hydrogen product stream 50 is split into three portions. A first portion of the hydrogen product stream is sent via line 52 to a dehydration unit 54 operating by glycol dehydration, which dries the product hydrogen stream, thereby producing a fuel-grade product hydrogen stream 56. A second portion of the hydrogen product stream is sent via line 58 to a pressure swing absorption unit 60, which purifies the hydrogen product by removing traces of methane and carbon oxides to provide a high purity hydrogen product stream 62. Pressure swing absorption unit 60 produces a tail gas stream 64 containing methane and carbon oxides as well as some hydrogen. A remaining third portion of hydrogen product stream 50 is combined with tail gas stream 64 to form a fuel gas stream 66 that is combusted with a combustion air stream 68 in a feed preheater and steam boiler unit 70 that is used to preheat the natural gas stream 10 and the natural gas and steam mixture 20 upstream of the synthesis gas production unit 22. Unit 70 also provides superheated medium pressure steam via line 18 to produce a flue gas stream 72.

[0065] In FIG. 2, a hydrogen product gas stream 50 is produced in the same manner as described in FIG. 1. In this embodiment, the hydrogen product stream 50 is split into two portions. A first portion of the hydrogen product stream is sent via line 52 to a dehydration unit 54 operating by glycol dehydration to dry the product hydrogen stream, thereby producing a fuel-grade product hydrogen stream 56. The remaining portion of the hydrogen product stream 50 is fed to a gas turbine unit 80 where it is combusted with air fed via line 82 to generate electricity 84 which is used in the process, for example to drive an air separation unit to provide an oxygen gas stream used in the autothermal reforming and oxidation vessels. Electricity may also be exported. Exhaust gas recovered from the gas turbine via line 86 is used to generate superheated medium pressure steam 18 in a heat recovery steam generator unit 88. Spent exhaust gas is recovered from the heat recovery unit 88 via line 90. In this embodiment, heating of the natural gas stream 10 prior to desulfurization is provided by heat exchange with the medium pressure steam produced in unit 88. Intermediate pressure steam 18 for the steam reforming is supplied by the same unit 88. However, heating of the natural gas and steam mixture 20 is provided by heat exchange in a feed-effluent heat exchanger (not shown) between the tube-side feed and shell-side effluent of the gas-heated reformer.

[0066] In Fig. 3, a natural gas stream 110 is processed in a purification and saturation unit 112 for desulfurization, including hydrodesulfurization over a hydrodesulfurization catalyst and adsorption of hydrogen sulfide over a hydrogen sulfide adsorbent, followed by saturation with steam using a saturator to form a saturated natural gas mixture 114. The steam ratio of the saturated natural gas mixture 114 can be increased by the addition of saturated medium pressure steam via line 130 and superheated medium pressure steam via line 176. The resulting gaseous mixture of steam and natural gas 120 is sent to a synthesis gas production unit 122, which includes a gas-heated reformer (GHR) and an autothermal reformer (ATR) in series. The gaseous mixture 120 is subjected to a first step of steam reforming in the gas-heated reformer over a granular nickel steam reforming catalyst arranged in a number of external heating tubes, followed by autothermal reforming with an oxygen gas stream 124 in the autothermal reformer containing a nickel steam reforming catalyst bed. The reformed gas recovered from the autothermal reformer is used to heat the tubes of the gas-heating reformer, thereby producing a partially cooled reformed gas 126, which is fed to a water-gas shift unit 128, which includes an isothermal water-gas shift vessel containing a bed of water-gas shift catalyst cooled by boiling water under pressure. The exothermic water-gas shift reaction generates steam, which is recovered from the water-gas shift unit 128 via line 130 and fed to the saturated natural gas stream 114. The hydrogen-rich gas mixture is recovered from the water-gas shift unit 128 via line 132 and fed to a heat recovery unit 134, where it is cooled below the dew point by heat exchange with cold water to condense water therefrom. Condensate is recovered from the cooled mixture using one or more gas-liquid separators (not shown) and fed to a saturator in unit 112. The dehydrated hydrogen-rich gas recovered from the heat recovery unit 134 is heated and fed via line 136 to an oxidation unit 138, which includes an oxidation vessel containing a fixed bed of promoted platinum CO oxidation catalyst. An additional oxygen-enriched gas stream is added to the oxidation vessel via line 140. Carbon monoxide is oxidized to form a carbon dioxide-enriched gas mixture. In this embodiment, the CO oxidation catalyst is operated adiabatically.The carbon dioxide-rich gas recovered from the oxidation unit 138 via line 142 is subjected to heat recovery in a low-temperature heat recovery unit 144, which cools the carbon dioxide-rich gas 142 below the dew point in heat exchange with cold water to condense the vapor. The heated water recovered from the unit 144 can be sent to a district heating plant 148 via line 146, and the cooled water can be returned to the heat recovery unit 144 via line 147. The liquid condensate can be separated from the carbon dioxide-rich gas in one or more gas-liquid separators (not shown) and fed to a saturator in the unit 112. The dehydrated carbon dioxide-rich gas recovered from the one or more gas-liquid separators is fed via line 150 to a carbon dioxide removal unit 152 operating by reactive absorption to remove carbon dioxide from the gas mixture, thereby producing a carbon dioxide stream 154 that is compressed and sent for storage or conversion to chemicals. By removing carbon dioxide from the feed gas 150, the carbon dioxide removal unit produces a hydrogen product gas stream that is recovered from the carbon dioxide removal unit via line 156. In this embodiment, all of the hydrogen product stream 156 is supplied to a gas turbine unit 158 ​​where it is combusted with air supplied via line 160 to generate electricity 162 which is used in the process, for example to drive an air separation unit to provide an oxygen gas stream used in the autothermal reforming and oxidation vessels. Electricity may also be exported. Exhaust gas recovered from the gas turbine via line 164 is used to generate and superheat steam in a heat recovery steam generator unit 166. Spent exhaust gas is recovered from the unit 166 via line 168. In this embodiment, superheated high pressure steam produced by the heat recovery steam generator in the unit 166 is supplied via line 170 to a steam turbine 172. The steam turbine 172 may be used to generate additional electricity 174 and provide superheated intermediate pressure steam via line 176 which is added to the saturated natural gas stream 114. The steam turbine exhaust stream 178 undergoes separate low temperature heat recovery in a heat recovery unit 180 and heated water is supplied from the unit 180 via line 182 to the district heating plant 148.The district heating plant returns cooled water to the low temperature heat recovery unit 180 via line 184. Recovered boiler feedwater from the low temperature heat recovery unit 180 is provided via line 186 and supplied to the heat recovery steam generator unit 166. Optionally, make-up boiler feedwater can be provided via line 188 to balance the steam extracted from the steam turbine 172 via line 176. In this embodiment, heating of the natural gas and steam mixture 120 is provided by heat exchange in a feed-effluent heat exchanger (not shown) between the tube-side feed and shell-side effluent of the gas-heated reformer.

[0067] The embodiment of Figure 1 can be used where at least a portion of the product hydrogen needs to meet high purity specifications, for example for use in a downstream synthesis process. In this case, the main benefit of the CO oxidation unit is to reduce on-site carbon emissions by reducing the amount of CO present in the fuel gas. The heat produced by the CO oxidation reaction is used to raise a portion of the process steam.

[0068] The embodiment of FIG. 2 may be used when all product hydrogen should be exported as fuel and high purity hydrogen is not required. In this case, one further advantage of the CO oxidation unit is that it eliminates the need for a PSA purification step. A further advantage of this is that all product hydrogen is available under pressure, since the low pressure PSA tail gas is also eliminated. Thus, a portion of the product hydrogen can be burned in a gas turbine without the need to recompress it. Since gas turbines generate mechanical work or electricity, this embodiment provides a more exergy-efficient way to utilize the high temperature heat generated by fuel combustion. Furthermore, decarbonized electricity is a welcome by-product and can contribute to meeting the ever-increasing demand for power generated by electric vehicles.

[0069] The embodiment of Figure 3 may be used when it is desirable to decarbonize the process for combined heat and power generation (CHP). In this case, a combined cycle power plant can be thermally integrated with a low carbon hydrogen plant, and low temperature waste heat from both plants can be exported to the heat network (district heating). In this embodiment, the CO oxidation unit is not operated isothermally to raise medium temperature steam. It is operated adiabatically at a lower temperature to increase the amount of low temperature heat that can be exported to the heat network.

[0070] FIG. 4 is a modification of the embodiment of FIG. 1 in which 100% of the hydrogen product gas stream is purified in a PSA unit.

[0071] In FIG. 4, a natural gas stream 200 is mixed with a small hydrogen stream provided via line 202 and fed via line 204 to a coil in a heater 206 where it is heated by burning a portion of the PSA off-gas to generate flue gas 210. The heated natural gas is fed from the heater 206 via line 212 to a purification unit 214 where it is subjected to hydrodesulfurization over a hydrodesulfurization catalyst followed by desulfurization by adsorption of hydrogen sulfide on a hydrogen sulfide adsorbent. The desulfurized natural gas is fed via line 216 to a saturator 218 where it is saturated with steam by contact with hot water fed to the saturator via line 220. A saturated natural gas stream is withdrawn from the saturator 218 via line 222. Medium pressure steam is added to the saturated natural gas via line 224 and superheated steam is added to the saturated natural gas via line 226. The resulting natural gas and steam mixture is fed via line 228 to a coil in heater 206 where it is heated before being fed via line 230 to externally heated catalyst-containing tubes of gas-heated reformer 232. The gaseous mixture 230 is subjected to a first step of steam reforming in gas-heated reformer 232. Partially reformed gas is withdrawn from gas-heated reformer 232 and fed via line 234 to a burner of autothermal reformer 236 where it is partially combusted with oxygen-containing gas fed via line 238 and passed through a fixed bed of steam reforming catalyst disposed downstream of the burner to further reform the gas mixture. The autothermal reformed gas mixture is withdrawn from autothermal reformer 236 and fed via line 240 to the shell side of gas-heated reformer 232 where it passes around catalyst-containing tubes, thereby providing heat for the first reforming step. Partially cooled reformed gas is withdrawn from the shell side of the gas-heated reformer via line 242, cooled in heat exchanger 246, and supplied via line 248 to an isothermal water-gas shift vessel 250 containing a bed of copper-based water-gas shift catalyst cooled by heat exchange against boiling water under pressure supplied to vessel 250 from steam drum 252 via line 254. An exothermic water-gas shift reaction occurs in vessel 250 generating steam which is returned to steam drum 252 via line 256.The hydrogen-rich gas mixture is withdrawn from the isothermal water-gas shift vessel 250 and fed, without cooling, via line 258 to the inlet of an oxidation vessel 260 containing a fixed bed of a precious metal oxidation catalyst cooled by heat exchange with boiling water under pressure, which is fed from the steam drum 252 to the oxidation vessel 260 via line 262. The oxidation vessel 260 is fed with an oxygen-containing gas via line 264. An exothermic oxidation reaction takes place, producing steam which is returned to the steam drum 252 via line 266. The pressure of the coolant for both vessels 250 and 260 is the same, which allows the use of a common steam drum 252. The intermediate pressure steam withdrawn from the steam drum 252 is fed to the saturated natural gas 222 via line 224. Carbon monoxide in the hydrogen-rich gas is oxidized in the oxidation vessel 260 to form carbon dioxide, and the carbon dioxide-rich gas is sent from the unit 260 to a heat recovery unit 270 via line 268. The carbon dioxide-rich gas is cooled in heat recovery unit 270 by heat exchange with water provided from saturator 218 via line 272. Heat recovery unit 270 produces a condensate stream which is withdrawn from heat recovery unit 270 via line 273. Cooled carbon dioxide-rich gas 276 from heat recovery unit 270 is further cooled below the dew point using cold water in heat exchanger 278 and the resulting mixture is fed to gas-liquid separator 280 where further liquid condensate is separated to form dehydrated carbon dioxide-rich gas. Condensate withdrawn from gas-liquid separator 280 via line 282 is combined with condensate stream 273 from heat recovery unit 270 and the combined stream is circulated via pump 284 to heat recovery unit 270 where it is heated against the carbon dioxide-rich gas. The resulting heated condensate is then mixed via line 288 with saturator water in line 272 which is fed to heat recovery unit 270 for further heating before being fed to saturator 218 via line 274, heat exchanger 246 and line 220. The dehydrated carbon dioxide enriched gas recovered from gas-liquid separator 280 is fed via line 290 to a carbon dioxide removal unit 292 operating by reactive absorption which removes carbon dioxide from the gas mixture, thereby producing a carbon dioxide stream 294 which is compressed and sent for storage or conversion to chemicals.By removing carbon dioxide from the feed gas 290, the carbon dioxide removal unit 292 produces a hydrogen product gas stream that is withdrawn via line 296. In this embodiment, all of the hydrogen product stream 296 is fed to the pressure swing absorption unit 298. The pressure swing absorption unit separates trace amounts of methane, carbon dioxide and nitrogen from the hydrogen product, thereby producing a purified hydrogen product stream, which is withdrawn from the pressure swing absorption unit 298 via line 300. The purified hydrogen product is then compressed in compressor 302 to provide a compressed purified hydrogen stream 304 for export. A small portion of the compressed purified hydrogen is fed from line 304 via line 202 to the natural gas stream 200 to provide hydrogen for hydrodesulfurization in the purification unit 214. The pressure swing absorption unit produces a tail gas stream 306 containing hydrogen and small amounts of methane and nitrogen, which is withdrawn via line 306 and used to heat the feed and generate steam for the process. A portion of the tail gas stream 306 is supplied via line 208 to the feed-fired heater 206 where it is combusted to generate heat. The remaining portion of the tail gas stream 306 is supplied via line 308 to fire a steam boiler and superheater 310 which produces superheated steam which is supplied from the boiler and superheater 310 to the saturated natural gas stream 222 via line 226. The boiler and superheater 310 produces flue gas 312.

[0072] The invention is further illustrated by the following calculated example of a process according to the flowsheet shown in FIG.

[0073] [Table 1]

[0074] [Table 2]

[0075] [Table 3]

[0076] [Table 4]

[0077] The examples demonstrate the low CO2 emissions achievable from the process.

Claims

1. 1. A process for producing hydrogen, comprising: (i) subjecting a gaseous mixture comprising hydrocarbons and steam to steam reforming in a gas-heated reformer or an adiabatic pre-reformer, followed by autothermal reforming with an oxygen-enriched gas in an autothermal reformer to produce a reformed gas mixture; (ii) increasing the hydrogen content of the reformed gas mixture by subjecting the reformed gas mixture to one or more water-gas shift stages in a water-gas shift unit to provide a hydrogen-enriched reformed gas; (iii) passing the hydrogen-rich reformed gas and the oxygen-rich gas through an oxidation unit containing an oxidation catalyst that converts carbon monoxide present in the hydrogen-rich reformed gas to carbon dioxide to form a carbon dioxide-rich gas mixture; (iv) cooling the carbon dioxide-enriched gas mixture and separating condensed water therefrom; (v) passing said carbon dioxide-enriched gas mixture to a carbon dioxide separation unit to provide a carbon dioxide gas stream and a hydrogen product gas stream.

2. 10. The process of claim 1, wherein the hydrocarbon is a methane-containing gas stream.

3. 3. The process of claim 1 or 2, wherein the hydrocarbon is desulfurized.

4. 3. The process of claim 1 or 2, wherein the steam to carbon ratio in the gaseous mixture is in the range of 0.9:1 to 5:

1.

5. 3. The process of claim 1 or 2, wherein the gaseous mixture comprising the hydrocarbon and steam is contacted with water in a saturator to form a saturated gas mixture.

6. 3. The process of claim 1 or 2, wherein the oxygen-enriched gas comprises at least 90% by volume O2.

7. 3. The process of claim 1 or 2, wherein the water-gas shift stage comprises an isothermal shift stage.

8. 3. The process of claim 1 or 2, wherein the oxidation catalyst comprises one or more of Pt, Pd, Rh, Ir, or Ru.

9. 3. The process of claim 1 or 2, wherein the carbon dioxide removal stage is preceded by one, two or three stages of cooling and separating process condensate from the carbon dioxide-rich gas mixture.

10. 3. The process of claim 1 or 2, wherein the carbon dioxide removal step is carried out using a physical scrubbing system or a reactive scrubbing system.

11. 3. The process of claim 1 or 2, wherein the carbon dioxide stream is compressed using an electrically driven compressor.

12. 3. The process of claim 1 or 2, wherein the hydrogen product is used in a downstream power process, a heating process, a downstream chemical synthesis process, or used to upgrade hydrocarbons.

13. 13. The process of claim 12, wherein the hydrogen product is purified in a purification unit to provide a purified hydrogen product and a tail gas.

14. 14. The process of claim 13, wherein a portion of the tail gas and a portion of the hydrogen product are supplied to one or more fired heaters used to heat feed steam or provide steam for the process, before or after purification in the purification unit.

15. 13. The process of claim 12, wherein at least a portion of the hydrogen product is combusted with air in a gas turbine to generate electricity and form hot exhaust gases.

16. 16. The process of claim 15, wherein the hot exhaust gas is used to heat the hydrocarbon feedstock and / or to generate steam for the steam reforming step.

17. 16. The process of claim 15, wherein the hot exhaust gas is used to heat steam that is fed to a steam turbine to generate additional electricity and produce steam for the steam reforming step.

18. 20. The process of claim 17, wherein one or more of the exhaust stream from the steam turbine, the hydrogen-enriched reformed gas mixture, and the carbon dioxide-enriched gas are used in a low temperature heat recovery process to provide heat for district heating.

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